Bubbling flow

Static mixer (bubble flow) 0.1-2 Up to 10 1-20 -1000 0.5 -Plug Plug 10-500... 

Bubble flow - The gas is roughly uniformly distributed in the form of small discrete bubbles in a continuous liquid phase. The flow pattern is designated as bubble flow (B) at low liquid flowrates, and as dispersed bubble (DB) at high liquid flow rates in which case the bubbles are finely dispersed in the liquid. 

The term three-phase fluidization requires some explanation, as it can be used to describe a variety of rather different operations. The three phases are gas, liquid and particulate solids, although other variations such as two immiscible liquids and particulate solids may exist in special applications. As in the case of a fixed-bed operation, both co-current and counter- current gas-liquid flow are permissible and, for each of these, both bubble flow, in which the liquid is the continuous phase and the gas dispersed, and trickle flow, in which the gas forms a continuous phase and the liquid is more or less dispersed, takes place. A well established device for countercurrent trickle flow, in which low-density solid spheres are fluidized by an upward current of gas and irrigated by a downward flow of liquid, is variously known as the turbulent bed, mobile bed and fluidized packing contactor, or the turbulent contact absorber when it is specifically used for gas absorption and/or dust removal. Still another variation is a three-phase spouted bed contactor. 

Cocurrent three-phase fluidization is commonly referred to as gas-liquid fluidization. Bubble flow, whether coeurrent or countereurrent, is eonveniently subdivided into two modes mainly liquid-supported solids, in which the liquid exeeeds the minimum liquid-fluidization veloeity, and bubble-supported solids, in whieh the liquid is below its minimum fluidization velocity or even stationary and serves mainly to transmit to the solids the momentum and potential energy of the gas bubbles, thus suspending the solids. 

Countereurrent bubble flow with liquid-supported solids, whieh ean be affeeted by downward liquid fluidization of partieles having a density lower than that of the liquid, has been referred to as inverse three-phase fluidization. The mass transfer potential of sueh a eountercurrent operation is worthy of study, especially for cases in whieh dispersion of the gas rather than the liquid is ealled for and the required gas-liquid ratio and throughput ean be effected without flooding. In contrast, the eorresponding eoeurrent mode has reeeived more attention than all other eases and eonstitutes the majority of the literature on three-phase fluidization. 

The exit region of a die used to blow plastic film is shown below. If the extruder output is 100 X 10 m /s of polythene at 170°C estimate the total pressure drop in the die between points A and C. Also calculate the dimensions of the plastic bubble produced. It may be assumed that there is no inflation or draw-down of the bubble. Flow data for polythene is given in Fig. 5.3. 

Aeeording to the penetration theory, liquid elements or pareels remain in eontaet with the gas for a limited time and subsequently are mixed with the bulk. This eoneept is partieularly suited to bubble flows beeause it refleets their... 

Slug Slugs of gas bubbles flowing through the liquid... 

On low gas content and low speed of gas-water mixture, the bubbles are comparatively small and distributed throughout the section of the pipe in an even manner this flow mode was termed emulsion or bubble flow ... 

An interesting and practically valuable result was obtained in [21] for PE + N2 melts, and in [43] for PS + N2 melts. The authors classified upper critical volumetric flow rate and pressure with reference to channel dimensions x Pfrerim y Qf"im-Depending on volume gas content

channel entrance (pressure of 1 stm., experimental temperature), x and y fall, in accordance with Eq. (24), to tp 0.85. At cp 0.80, in a very narrow interval of gas concentrations, x and y fall by several orders. The area of bubble flow is removed entirely. It appears that at this concentration of free gas, a phase reversal takes place as the polymer melt ceases to be a continuous phase (fails to form a continuous cluster , in flow theory terminology). The theoretical value of the critical concentration at which the continuous cluster is formed equals 16 vol. % (cf., for instance, Table 9.1 in [79] and [80]). An important practical conclusion ensues it is impossible to obtain extrudate with over 80 % of cells without special techniques. In other words, technology should be based on a volume con-... 

In high heat flux (heat transfer rate per unit area) boilers, such as power water tube (WT) boilers, the continued and more rapid convection of a steam bubble-water mixture away from the source of heat (bubbly flow), results in a gradual thinning of the water film at the heat-transfer surface. A point is eventually reached at which most of the flow is principally steam (but still contains entrained water droplets) and surface evaporation occurs. Flow patterns include intermediate flow (churn flow), annular flow, and mist flow (droplet flow). These various steam flow patterns are forms of convective boiling. 

In the first class, the particles form a fixed bed, and the fluid phases may be in either cocurrent or countercurrent flow. Two different flow patterns are of interest, trickle flow and bubble flow. In trickle-flow reactors, the liquid flows as a film over the particle surface, and the gas forms a continuous phase. In bubble-flow reactors, the liquid holdup is higher, and the gas forms a discontinuous, bubbling phase. 

Two types of fixed-bed operations, characterized by distinctly different flow patterns, are in current industrial use. These are usually described as trickle-flow operation and bubble-flow operation. In both cases, a lower limit exists for the particle size, usually about k in. 

In bubble-flow operation, the gaseous phase moves upwards as discrete bubbles. The liquid phase may be in either co- or countercurrent flow. The liquid holdup is relatively high. 

It may be noted that trickle-flow operations is not always clearly distinguished in the literature from fixed-bed bubble-flow operation. The two... 

Zabor et al. (Zl) have described studies of the catalytic hydration of propylene under such conditions (temperature 279°C, pressure 3675 psig) that both liquid and vapor phases are present in the packed catalyst bed. Conversions are reported for cocurrent upflow and cocurrent downflow, it being assumed in that paper that the former mode corresponds to bubble flow and the latter to trickle-flow conditions. Trickle flow resulted in the higher conversions, and conversion was influenced by changes in bed height (for unchanged space velocity), in contrast to the case for bubble-flow operation. The differences are assumed to be effects of mass transfer or liquid distribution. 

Fixed-bed bubble-flow operation has been the subject of considerably less experimental work than has trickle-flow operation. This reflects the fact that bubble-flow operation has been of much more limited industrial importance. 

Some data on gas holdup are also reported by Stemerding (SI6). Hoogendoorn and Lips (H10) have reported gas-holdup data for counter-current bubble flow in the experimental system described in Section V,A,4. Gas holdup was not influenced by changes of liquid flow rate, but increased with nominal gas velocity in the range from 0.03 ft/sec to 0.3 ft/sec. The results are somewhat lower than those obtained by Weber, the difference being explained as due to the difference in gas distributor. Weber used a porous plate and Hoogendoorn and Lips a set of parallel nozzles. 

Schoenemann (S4) reported qualitatively that the liquid residence-time distribution for cocurrent upward bubble flow was narrower than that observed in trickle-flow operation. 

No data on heat transfer in fixed-bed bubble-flow operation appear to have been published. 

The only data on chemical conversion in fixed-bed bubble-flow operation that have come to the author s attention are the few results referred to in Section V,A,6 obtained by Zabor et al. (Zl) for catalytic mixed-phase hydration of propylene. 

A theory of two-phase bubble flow has been developed by Nicklin (N1), who shows that the motion of bubbles arises partly from buoyancy and partly from the nominal velocity caused by the entry of the two phases into the tube. Theoretical and semiempirical studies of bubble flow have also been presented by Azbel (A2) and by Azizyan and Smirnov (A3), and further experimental data on holdup have been recently reported by Yoshida and Akita (Yl), by Braulick et al. (B9) and by Towell et al. (T3). 

Gal-Or and Resnick (Gl) have developed a simplified theoretical model for the calculation of mass-transfer rates for a sparingly soluble gas in an agtitated gas-liquid contactor. The model is based on the average gas residencetime, and its use requires, among other things, knowledge of bubble diameter. In a related study (G2) a photographic technique for the determination of bubble flow patterns and of the relative velocity between bubbles and liquid is described. 

The liquid residence-time distribution is close to plug flow in trickle-flow operation and corresponds to perfect mixing in the stirred-slurry operation, whereas the other types of bubble-flow operation are characterized by residence-time distributions between these extremes. 

The flow patterns of agitated liquid have been studied extensively (Al, B11, F6, K5, M6, N2, R12, V5), usually by photographic methods. Apparently no work has been reported on bubble-flow patterns and relative velocities in agitated gas-liquid dispersions. Some simple pictures have been presented that only show the same details that may be seen with the unaided eye (Bll, F6, Y4). 

A Reynolds number for the bubble flow which represents the term for agitation may be defined as ... 

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